Process for improving the cetane number of a gas oil cut

ABSTRACT

A process for transforming a gas oil cut from a conversion process or from an aromatic crude is described, the aim of the process being to improve the cetane number of said cut. The process comprises at least one hydrogenation step in which said gas oil cut is passed, in the presence of hydrogen, over a catalyst comprising an amorphous mineral support, at least one compound of a group VIB metal, at least one compound of a non noble group VIII metal and at least phosphorous or a compound of phosphorous, the process then comprising a hydrocracking step in which the hydrogenated feed is passed, in the presence of hydrogen, over a catalyst comprising an acidic support, at least one compound of a group VIB metal and at least one compound of a non noble group VIII metal.

The present invention relates to the field of fuels for internalcombustion engines. More particularly, it relates to the production of afuel for a compression ignition engine, and to the fuel obtainedtherefrom.

Whether from straight run distillation of a crude petroleum or from aconversion process such as catalytic cracking, gas oil cuts stillcontain non negligible quantities of aromatic compounds,nitrogen-containing compounds and sulphur-containing compounds. Currentlegislation in the majority of industrialised countries requires that afuel for use in engines must contain less than 500 parts per million(ppm) of sulphur. In the vast majority of those countries, there arecurrently no regulations imposing a maximum aromatic compound andnitrogen content. However, a number of countries or states, for exampleSweden and California, envisage limiting the quantity of aromaticcompounds to a value of less than 20% by weight, or even to less than10%, and some experts think that that limit should be 5%. In Sweden inparticular, some classes of diesel fuel already have to satisfy verystrict regulations. Thus in that country, class II diesel fuel must notcontain more than 50 ppm of sulphur and more than 10% by weight ofaromatic compounds, and class I fuel must not contain more than 10 ppmof sulphur and 5% by weight of aromatic compounds. Currently in Sweden,class III diesel fuel must contain less than 500 ppm of sulphur and lessthan 25% by weight of aromatic compounds. Similar limits have to besatisfied to sell that type of fuel in California.

Meanwhile, motorists in a number of countries are pressing forlegislation to require oilmen to produce and sell a fuel with a minimumcetane number of constantly improving quality. Current Europeanlegislation requires a minimum cetane number of 49 which from the year2000 will rise to 51, probably at least 53 and more probably in therange 55 to 70.

Further, the same European regulations predict a tightening of theregulations regarding the density, the 95% point, sulphur content andpolyaromatics content.

A number of specialists seriously envisage the possibility of a futurestandard imposing a nitrogen content of less than about 200 ppm, forexample, and even less than 100 ppm by weight. A low nitrogen contentresults in a better product stability and is generally desired both bythe vendor of the product and by the manufacturer.

It is thus necessary to develop a reliable and effective process whichenables a product to be produced which has improved characteristics bothas regards the cetane number and the aromatic compound content, sulphurcontent and nitrogen content. The gas oil cuts originate either fromstraight run crude oil distillation, or from catalytic cracking: i.e.,light distillate cuts (LCO, Light Cycle Oil), heavy fraction cuts (HCO,Heavy Cycle Oil), or from another conversion process (cokefaction,visbreaking, residue hydroconversion, etc.), or from gas oils from thedistillation of aromatic or naphthenoaromatic Hamaca, Zuata, or El Paotype crude oil. The production of an effluent which is directly andintegrally upgradeable as a very high quality fuel cut is particularlyimportant.

Conventional processes can improve the cetane number to an extent whichsatisfies current cetane number regulations for the majority of feeds.However, with gas oil cuts originating from a catalytic cracking typeconversion process or in the case of particularly severe specifications,this increase reaches a limit which cannot be exceeded using theconventional sequences of such processes.

Further, a well known advantage of these catalysts is that a prolongedservice life is possible without observing any deactivation.

The prior art describes processes for hydrogenating petroleum cuts whichare particularly rich in aromatic compounds which use a catalyst, forexample United States patent U.S. Pat. No. 5,037,532 or the publication“Proceedings of the 14^(th) World Petroleum Congress, 1994, p. 19-26”which describe processes which lead to the production of hydrocarboncuts, and increase in the cetane number is obtained by intensehydrogenation of the aromatic compounds.

We have now sought to produce fuels with a cetane number of the sameorder as those obtained by conventional hydrogenation processes orhigher but without having recourse to hydrogenation which is toointense.

The present invention is distinguished over the prior art in that itcombines hydrocracking with hydrogenation.

Such a combination has already been described for the treatment of heavyfeeds, for example in French patent FR-A-2 600 669.

In that patent, the treated feed contains at least 50% by weight ofconstituents boiling above 375° C. and the aim of the process is toconvert at least 70% by volume of those heavy constituents toconstituents with a boiling point of less than 375° C.

At the end of the process, at least one cut is produced with a boilingpoint below 375° C. (gasoline, gas oil) and a heavy cut is produced witha boiling point of at least 375° C. which can be recycled to improveconversion. The light compounds are, of course, separated out (residualH₂, C₁-C₄, H₂S, NH₃ . . . ).

Thus this process comprising a hydrotreatment step followed by ahydrocracking step uses a zeolitic catalyst converts a heavy cut to agas oil (250-375° C.) and a gasoline (150-250° C.) with the highestyield possible.

The Applicant has been able to establish that, compared with the priorart hydrogenation to treat gas oil cuts, the process of the invention,combining hydrogenation and hydrocracking, breaks the conventionalcetane limits encountered in conventional hydrogenation processes andmore substantially reduces the 95% ASTM point (the point correspondingto the boiling point of 95% of the cut).

More precisely, the invention provides a process for converting a gasoil cut into a high cetane number fuel which is dearomatised,desulphurised and has good cold properties, the process comprising thefollowing steps:

-   a) at least one first step termed hydrogenation in which said gas    oil cut is passed, in the presence of hydrogen, over a catalyst    comprising an amorphous mineral support, at least one metal or    compound of a metal from group VIB of the periodic table (Handbook    of Chemistry and Physics, 76^(th) Edition, 1995-1996) in a quantity,    expressed as the weight of metal with respect to the weight of    finished catalyst, of about 0.5% to 40%, at least one non noble    metal or compound of an non noble metal from group VIII of the    periodic table in a quantity, expressed as the weight of metal with    respect to the weight of finished catalyst, of about 0.01% to 30%,    and of phosphorous or at least one compound of phosphorous in a    quantity, expressed as the weight of phosphorous pentoxide with    respect to the weight of the support, of about 0.001% to 20%; and-   b) at least one second step, termed hydrocracking, in which the    hydrogenated product from the first step is passed, in the presence    of hydrogen, over a catalyst comprising a mineral support which is    partly zeolitic, at least one metal or compound of a metal from    group VIB of the periodic table in a quantity, expressed as the    weight of metal with respect to the weight of finished catalyst, of    about 0.5% to 40% and at least one non noble metal or compound of a    non noble metal from group VIII in a quantity, expressed as the    weight of metal with respect to the weight of finished catalyst, of    about 0.01% to 20%, the light compounds then being separated from    the hydrocracking effluent.

This two-step process essentially comprises substantial or managedhydrogenation of the aromatic compounds—depending on the amount ofaromatic compounds which are to be in the final product—, thenhydrocracking intended to open the naphthenes produced in the firststep, to form paraffins.

These feeds are treated in hydrogen in the presence of catalysts, thistreatment enabling the aromatic compounds present in the feed to behydrogenated; it can also simultaneously carry out hydrodesulphurisationand hydrodenitrogenation.

In the process of the present invention, the operating conditions forhydrogenation (or hydrotreatment) are as follows: the hourly spacevelocity (HSV) is in the range 0.1 to 30 volumes of liquid feed pervolume of catalyst per hour, preferably in the range 0.2 to 10; thetemperature at the reactor inlet is in the range 250° C. to 450° C.,preferably in the range 320° C. to 400° C.; the reactor pressure is inthe range 0.5 to 20 MPa, preferably in the range 4 to 15 MPa; the purehydrogen recycle rate is in the range 100 to 2500 Nm³/m³ of feed,preferably in the range 200 to 2100 Nm³/m³, more advantageously lessthan 2000 Nm³/m³. The hydrogen consumption in the process can be up toabout 5% by weight of the feed (0.5-4.5% in general).

The hydrogenation catalyst comprises, on an amorphous mineral support,at least one metal or compound of a metal from group VIB of the periodictable, such as molybdenum or tungsten, in a quantity, expressed as theweight of metal with respect to the weight of finished catalyst, in therange 0.5% to 40%, preferably in the range 2% to 30%, at least one nonnoble metal or a compound of a non noble metal from group VIII of saidperiodic table, such as nickel, cobalt or iron, in a quantity, expressedas the weight of metal with respect to the weight of finished catalyst,in the range 0.01% to 30%, preferably in the range 0.1% to 10%,phosphorous or at least one phosphorous compound, in a quantity,expressed as the weight of phosphorous pentoxide with respect to theweight of the support, in the range 0.001% to 20%. The catalyst can alsocontain boron or at least one compound of boron in a quantity, expressedas the weight of boron trioxide with respect to the weight of thesupport, in the range 0.001% to 10%. The amorphous mineral support is,for example, alumina or silica-alumina. In a particular embodiment ofthe invention, cubic gamma alumina is used which preferably has aspecific surface area of about 50 to 500 m²/g.

The hydrogenation catalyst used in the present invention preferablyundergoes a sulphurisation treatment to at least partially transform themetallic species to the sulphide before bringing them into contact withthe feed to be treated. This sulphurisation activation treatment is wellknown to the skilled person and can be carried out using any methodwhich is already known in the literature.

One conventional method which is well known to the skilled personconsists of heating the catalyst in the presence of hydrogen sulphide orof a hydrogen sulphide precursor to a temperature in the range 150° C.to 800° C., preferably in the range 250° C. to 600° C., generally in atraversed bed reaction zone.

The term “hydrogen sulphide precursor” as used in the presentdescription means any compound which can react under the operatingconditions of the reaction to give hydrogen sulphide.

The hydrogenated products from the first step may or may not undergo atreatment selected from the group formed by gas-liquid separations anddistillations. The liquid phase then undergoes hydrocracking in step b)of the present invention.

In the process of the present invention, the operating conditions forthe hydrocracking step are as follows: the hourly space velocity (HSV)is about 0.1 to 30 volumes of liquid feed per volume of catalyst perhour, preferably in the range 0.2 to 10, the reactor inlet temperatureis in the range 250° C. to 450° C., preferably in the range 300° C. to400° C.; the reactor pressure is in the range 0.5 to 20 MPa, preferablyin the range 4 to 15 MPa and more preferably in the range 7 to 15 MPa;the pure hydrogen recycle rate is in the range 100 to 2200 Nm³/m³ offeed. Under these conditions, conversion is regulated as a function ofthe cetane number and the other properties (density, T95 . . . ) to beobtained. The total conversion (hydrocracking b)+ that obtained duringhydrogenation step a)) can be higher than 50% or less than 50% (5-50%,for example) depending on the cut to be treated.

The catalyst of the second step generally comprises at least onezeolite, at least one support and at least one hydro-dehydrogenatingfunction.

An acidic zeolite is particularly advantageous in this type ofembodiment, for example a faujasite type zeolite, preferably a Yzeolite. The zeolite weight content is in the range 0.5% to 80%,preferably in the range 3% to 50% with respect to the finished catalyst.Advantageously, a Y zeolite with a lattice parameter of 24.14×10⁻¹⁰ m to24.55×10⁻¹⁰ m is used.

The hydro-dehydrogenating function of the catalyst can advantageously beprovided by a combination of metals: further, the catalyst contains atleast one oxide or sulphide of a group VIB metal such as molybdenum ortungsten in a quantity, expressed as the weight of metal with respect tothe weight of finished catalyst, in the range 0.5% to 40%, and at leastone non noble metal or a compound of a non noble metal from group VIII,such as nickel, cobalt or iron in a quantity, is expressed as the weightof metal with respect to the weight of finished catalyst, in the range0.01% to 20%, preferably in the range 0.1% to 10%. These metals aredeposited on a support selected from the group formed by alumina,silica, silica-alumina, boron oxide, magnesia, silica-magnesia,zirconia, titanium oxide, clay, used alone or as a mixture, the supportrepresenting the complement to 100% of the other constituents of thecatalyst. The hydrocracking catalyst used in the present inventionpreferably undergoes a sulphurisation treatment to transform at least aportion of the metallic species to sulphides before bringing them intocontact with the feed to be treated. This sulphurisation activationtreatment is well known to the skilled person and can be carried outusing any method which is already known in the literature.

One conventional method which is well known to the skilled personconsists of heating the 25 catalyst in the presence of hydrogen sulphideor of a hydrogen sulphide precursor to a temperature in the range 150°C. to 800° C., preferably in the range 250° C. to 600° C., generally ina traversed bed reaction zone.

U.S. Pat No. 5,525,209 characterizes a particularly advantageous HY acidzeolite by different specifications: a SiO₂/Al₂O₃ mole ratio in therange 8 to 70, preferably in the range 12 to 40; a sodium content ofless than 0.15% by weight determined for the zeolite calcined at 1100°C.; a lattice parameter “a” of the unit cell in the range 24.55×10⁻¹⁰ mto 24.24×10⁻¹⁰ m, preferably in the range 24.38×10−10 m to 24.26×10⁻¹⁰m; a sodium ion take-up capacity C_(Na), expressed in grams of Na per100 grams of modified zeolite, neutralised then calcined, of over 0.85;a specific surface area, determined by the BET method, of more thanabout 400 m²/g, preferably more than 550m²/g; a water vapour adsorptioncapacity for a partial pressure of 2.6 torrs (34.6 MPa) of more thanabout 6% at 25° C.; a pore distribution in the range 1% to 20%,preferably in the range 3% to 15%, of the pore volume contained in poreswith a diameter in the range 20×10⁻¹⁰ m to 80×10⁻¹⁰ m; the remainder ofthe pore volume mainly being contained in pores with a diameter of lessthan 20×10⁻¹⁰ m.

In general, the Y—Na zeolite from which the HY zeolite is prepared has aSiO₂/Al₂O₃ mole ratio in the range 4 to 6; it is appropriate to firstreduce the amount of sodium (by weight) to a value of the order of 1% to3%, preferably to less than 2.5%; the Y—Na zeolite also generally has aspecific surface area in the range about 750 m²/g to 950 m²/g.

A number of variations of the preparations exist in which thehydrothermal treatment of the zeolite is generally followed by an acidtreatment.

The effluent obtained from hydrocracking is fractionated to separate thelight (cracked) products, i.e., products boiling below 150° C. ingeneral, or below 180° C. or another temperature selected by therefiner. Thus at least one 150° C.+ or 180° C.+ gas oil cut is obtained.If the feeds contain compounds with a boiling point of more than 370°C., they can advantageously be separated, preferably to recycle them tothe hydrogenation and/or hydrocracking step. Instead of cutting them at370° C., they can be cut at a lower temperature, for example at 350° C.,depending on the refiner's requirements.

The present invention thus enables gas oils to be obtained with a cetanenumber, and possibly the aromatic compound content, which is improvedsuch that the cuts can satisfy the current and future regulations. Thesegas oil cuts can be sold directly.

The present invention can maximally upgrade all of the productscontained in the treated petroleum cut. The yield of upgradeableproducts is close to 99% of the amount of hydrocarbons; in contrast toconventional processes, there are no liquid or solid waste products tobe incinerated.

The gas oil feeds to be treated are preferably light gas oils such asstraight run gas oils, gas oils from fluid catalytic cracking (FCC) orLCO. They generally have an initial boiling point of at least 180° C.and a final boiling point of at most 370° C. More broadly, the inventioncan be applied to gas oil cuts with an initial boiling point of at least150° C., at least 80% by weight of which boils at at most 370° C., andadvantageously at least 90% of which boils at at most 370° C. Thecomposition by weight per hydrocarbon family of these feeds variesdepending on the ranges. In a typical composition, the contents (byweight) of paraffins are in the range 5.0% to 30.0%, of naphthenes inthe range 5.0% to 40.0% by weight and of aromatic compounds in the range40.0% to 80.0%. Less aromatic feeds containing less than 40% ofaromatics and generally 20% to less than 40% of aromatics can also betreated, the naphthene content possibly rising to 60%.

The following examples illustrate the invention without limiting itsscope.

In the Examples below, the catalyst used in the hydrogenation step hadthe following characteristics: the nickel content, in the oxide form,was 3%; the molybdenum content, in the oxide form, was 16.5% with 6% ofphosphorous pentoxide on alumina. For hydrocracking, a catalyst wasadvantageously used in which the support was alumina. This catalystcontained 12% by weight of molybdenum, 4% of nickel in the form ofoxides and 10% of Y zeolite, this catalyst being described in Example 2of U.S. Pat. No. 5,525,209.

These catalysts were sulphurised using a mixture ofn-hexane/DMDS+aniline up to 320° C. after 3000 hours of continuousoperation, no deactivation of the catalysts as described in the examplewas observed.

EXAMPLE 1

The feed was treated in a pilot unit comprising two reactors in seriesunder the following conditions: the space velocity in the two reactorswas 0.29 volumes of liquid feed per volume of catalyst per hour, thetemperature at the first reactor inlet was 380° C. for hydrogenation and390° C. for hydrocracking; the pressure in the two reactor was 14 MPa.In each reactor, the hydrogen recycle rate was 2000 Nm³ per m³ of feed.The characteristics of the feed and the 190° C.+ product obtained aftereach step are shown in Table 1, after the hydrocracking step and afterdistillation. TABLE 1 190° C. + product Characteristics Feed afterhydrocracking Density at 15° C. 0.947 0.831 Pour point ° C. 3 −7 Motorcetane number 32 56 Total nitrogen (by weight) ppm 1290 <1 Sulphur (byweight) ppm 19700 <1 Paraffins % (by weight) 15 30 Naphthenes % (byweight) 17.3 69 Aromatic compounds % (by weight) 67.7 1 H₂ consumption %(by weight) 4.22 T95 ° C. 397 353

EXAMPLE 2

The feed was treated in a pilot unit comprising two reactors in seriesunder the following conditions: the space velocity in the two reactorswas 0.25 volumes of liquid feed per volume of catalyst per hour, thetemperature at the first reactor inlet was 385° C. for hydrogenation andin the second reactor, it was 375° C. for hydrocracking; the pressure inthe two reactors was 14 MPa. In each reactor, the hydrogen recycle ratewas 2000 Nm³ per m³ of feed. The characteristics of the feed and theproducts obtained after each step are shown in Table 2. TABLE 2Characteristics Feed Product after hydrocracking Density at 15° C. 0.9510.827 Pour point ° C. −36 −45 Motor cetane 18 53 Total nitrogen (byweight) ppm 826 <1 Sulphur (by weight) ppm 17600 <1

EXAMPLE 3

The feed was treated in the pilot unit of Example 1 comprising tworeactors in series, under the following conditions: the space velocityin the two reactors was 0.25 volumes of liquid feed per volume ofcatalyst per hour, the temperature at the first reactor inlet was 360°C. for hydrogenation and in the second reactor, it was 367° C. forhydrocracking; the pressure in the two reactors was 14 MPa. In eachreactor, the hydrogen recycle rate was 2000 Nm³ per m³ of feed. Thecharacteristics of the feed and the products obtained after each stepare shown in Table 3. TABLE 3 Product 150° C. + after product hydro-after Characteristics Feed genation hydrocracking Density at 15° C.0.951 0.874 0.835 Motor cetane number 18 33 44 Total nitrogen (byweight) ppm 830 <1 <1 Sulphur (by weight) ppm 17600 <30 <30 Paraffins %(by weight) 11 8 11 Naphthenes % (by weight) 10 87 85 Aromatic compounds% 79 5 4 (by weight) H₂ consumption % (by weight) 3.26 4.73 95% TBPpoint ° C. 378 342 322

It can be seen that, using the process of the invention (Examples 1 and2), with feeds with a high aromatic compound content, a final product isobtained which has the following characteristics: a high cetane number,low aromatic compound contents, in particular di- and polyaromaticcompounds, low sulphur, low nitrogen, a low pour point and a low 95%point. The gas oil cut obtained by this process was very high quality,it satisfied the regulations, even the most draconian thereof, imposedby different states.

This Example 3 shows the gain of the hydrocracking step in the qualityof the products; the gains obtained on the single hydrocracking catalystwere 39/1000ths of density, 22° C. at the 95% point and 11 cetanepoints.

This process for improving the cetane number in two steps produces a gasoil cut with a high cetane number. Thus the base cut can be hydrogenatedto a greater or lesser extent depending on whether the regulations forthe aromatic compounds of a given country are to be satisfied, but inall cases, hydrogen is saved compared with conventional processes forimproving gas oil cuts.

The invention has two major advantages: it can economise on hydrogensince a less intense hydrogenation is carried out to obtain the samecetane number; it can also enable a reserve of aromatic compounds to beconstituted which can, as required, be hydrogenated in a subsequenthydrogenation step, which means a potential increase in the cetanenumber. The latter case more particularly concerns starting gas oil cutswith high aromatic compound contents (40-80% by weight). Thehydrogenation step is carried out with any known hydrogenation catalyst,in particular those containing at least one noble metal deposited on anamorphous refractory oxide support (for example alumina). A preferredcatalyst contains at least one noble metal (preferably platinum), atleast one halogen (preferably 2 halogens: chlorine and fluorine) and amatrix (preferably alumina). The hydrogenation step can be carried outon the total effluent leaving the hydrocracking step, separation of the150− compounds (preferably 180− compounds) thus taking place after thishydrogenation step. The hydrogenation step can also be carried out onthe 150+ cut (or 180+ cut depending on the fraction selected),optionally followed by separation of the 150− (or 180−) compounds.

The limit imposed by conventional intense hydrogenation processes isfixed by the amount of aromatic compounds. Once these aromatic compoundshave all been hydrogenated, there is no possibility of increasing thecetane number, but in contrast by combining hydrocracking withhydrogenation, the cetane number can be increased still further, byincreasing the paraffin content in the cut. When gas oil cuts with lowaromatic compound contents (20% to less than 40%) are used, thecombination of the invention of the hydrogenation step then thehydrocracking step can produce a high cetane number, which could not beobtained with intense hydrogenation used in the prior art. Thus thesequence of processes we propose here enables the limit imposed byintense hydrogenation processes to be broken and the cetane number canbe increased beyond regulation requirements.

With the process of the invention, fuels with sulphur contents below 500ppm, or even below 50 ppm or even below 10 ppm are produced. At the sametime, the cetane numbers remain at least 49 or at least 50. The aromaticcompound content is generally at most 20% (5-20%) and the polyaromaticcompound content is reduced to below 1%.

Compared with the conventional intense hydrogenation process, theprocess of the invention can produce larger gains in the propertieslisted below. The gain is the difference observed between the values ofthe property for the product and that for the starting product. Densityat 15° C. Gain generally about 100/1000ths and more Cetane (150+ cut)Gain of at least 20 or 25 which can rise to +35 or more as opposed toabout 20 in hydrogenation processes 95% point Gains of 25° C. to 60° C.or more, as opposed to 10-20° C. maximum for hydrogenation.

These values are given by way of indication only, and do not constitutea minimum to achieve nor a maximum achieved.

1-14. (Cancelled).
 15. A fuel obtained using a process for converting agas oil cut into a high cetane number fuel which is dearomatised anddesulphurised, said process comprising at least one first step (a) inwhich hydrogenation is performed to produce a hydrogenated product,wherein said gas oil cut is passed, in the presence of hydrogen, over acatalyst comprising an amorphous mineral support, at least one metal orcompound of a metal from group VIB of the periodic table in a quantity,expressed as the weight of metal with respect to the weight of finishedcatalyst, of about 0.5% to 40%, at least one non noble metal or compoundof a non noble metal from group VIII of the periodic table in aquantity, expressed as the weight of metal with respect to the weight offinished catalyst, of about 0.01% to 30%, and phosphorous or at leastone phosphorous compound in a quantity, expressed as the weight ofphosphorous pentoxide with respect to the weight of the support, ofabout 0.001% to 20%; and at least one second step (b) in whichhydrocracking is performed to produce an effluent containing lightcompounds, wherein said hydrogenated product from the first step ispassed, in the presence of hydrogen, over a catalyst comprising amineral support which is partly zeolitic, at least one metal or compoundof a metal from group VIB of the periodic table in a quantity, expressedas the weight of metal with respect to the weight of finished catalyst,of about 0.5% to 40% and at least one non noble metal or compound of anon noble metal from group VIII in a quantity, expressed as the weightof metal with respect to the weight of finished catalyst, of about 0.02%to 20%, and then separating said light compounds from the hydrocrackingeffluent, wherein said gas oil cut introduced into the hydrogenationstep has an initial boiling point of at least 150° C., at least 90 wt %of which boils at at most 370° C., wherein total conversion of the gasoil cut in the hydrocracking and hydrogenation steps is less than 50%w/w.
 16. A fuel according to claim 15, wherein the boiling point of thegas oil cut is in the range 180-370° C.
 17. A fuel according to claim15, wherein the gas oil cut has an aromatic compound content of at least20% by weight and less than 40% by weight.
 18. A fuel according to claim15, wherein the metal from group VIB of the catalyst of step b) ismolybdenum or tungsten and the metal from group VIII of the catalyst ofstep b) is nickel, cobalt or iron.
 19. A fuel according to claim 15,wherein the operating conditions for steps a) and b) comprise atemperature of about 250C to about 450C, a total pressure of about 0.5to 20 MPa and a global hourly space velocity of liquid feed of about 0.1to about 30 h⁻¹.
 20. A fuel according to claim 15, wherein the catalystof step a) comprises boron or at least one compound of boron.
 21. A fuelaccording to claim 15, wherein said process further comprises subjectingeffluent from the hydrocracking step to a hydrogenation step.
 22. A fuelobtained by a process for converting a gas oil cut into a high cetanenumber fuel which is dearomatised and desulphurised, said processcomprising at least one first step (a) in which hydrogenation isperformed to produce a hydrogenated product, wherein said gas oil cut ispassed, in the presence of hydrogen, over a catalyst comprising anamorphous mineral support, at least one metal or compound of a metalfrom group VIB of the periodic table in a quantity, expressed as theweight of metal with respect to the weight of finished catalyst, ofabout 0.5% to 40%, at least one non noble metal or compound of a nonnoble metal from group VIII of the periodic table in a quantity,expressed as the weight of metal with respect to the finished catalyst,of about 0.01% to 30%, and phosphorous or at least one phosphorouscompound in a quantity, expressed as the weight of phosphorous pentoxidewith respect to the weight of the support, of about 0.001% to 20%; andat least one second step (b) in which hydrocracking is performed toproduce an effluent containing light compounds, wherein saidhydrogenated product from the first step is passed, in the presence ofhydrogen, over a catalyst comprising a mineral support which is partlyzeolitic, at least one metal or compound of a metal group VIB of theperiodic table in a quantity, expressed as the weight of metal withrespect to the weight of finished catalyst, of about 0.5% to 40% and atleast one non noble metal or compound of a non noble metal from groupVIII in a quantity, expressed as the weight of metal with respect to theweight of finished catalyst, of about 0.02% to 20%, and then separatingsaid light compounds from the hydrocracking effluent, wherein said gasoil cut introduced into the hydrogenation step has an initial boilingpoint of at least 150° C., at least 90 wt % of which boils at at most370° C.